Production of high purity and ultra-high purity gas

ABSTRACT

Trace amounts of carbon monoxide and optionally hydrogen are removed from gaseous feed streams by passing the feed stream through a carbon monoxide adsorbent ( 33 ) prior to passing it through a supported metal catalyst ( 34 ). The invention saves significant capital and operational costs over existing processes.

PRIORITY

This application claims the benefit of U.S. provisional applicationsSer. No. 60/384,612, filed May 31, 2002 and Ser. No. 60/385,148, filedJun. 3, 2002. The contents of each application are herein incorporatedby reference.

FIELD OF THE INVENTION

This invention relates to the removal of impurities from feed airstreams and more particularly to the removal of hydrogen (H₂) and carbonmonoxide (CO) from feed air streams.

BACKGROUND OF THE INVENTION

Carbon monoxide (CO) and hydrogen (H₂) can be present in air atconcentrations of up to about 50 ppm and 10 ppm respectively, althoughtypical concentrations in air are on the order of 1 ppm CO and 1 ppm H₂.Normal cryogenic distillation processes used to produce ultra highpurity (UHP) nitrogen (N₂) do not remove hydrogen and remove only asmall portion of the CO. Unless removed by alternative means, thesemolecules will contaminate the product nitrogen at a concentration up toabout two and a half times their concentration in the feed air. Sincethe electronic industry demands very high purity nitrogen product(typically having on the order of 5 ppb CO or less and 5 ppb H₂ orless), CO and H₂ have to be removed from feed air.

Air also contains other contaminants such as water (H₂O), carbon dioxide(CO₂) and hydrocarbons. In cold sections of the distillation separationprocess (such as heat exchangers and separation columns), water and CO₂can solidify and block the heat exchangers or other parts in thedistillation columns. Acetylene and other hydrocarbons in air alsopresent a potential problem because they can accumulate in the liquidoxygen (O₂) and create an explosion hazard. It is therefore desirable toremove these impurities prior to the cryogenic distillation of air.

Air prepurification can be accomplished using pressure swing adsorption(PSA), temperature swing adsorption (TSA) or a combination of both(TSA/PSA) incorporating either a single adsorbent or multipleadsorbents. When more than one adsorbent is used, the adsorbents may beconfigured as discrete layers, as mixtures, composites or combinationsof these. Impurities such as H₂O and CO₂ are commonly removed from airusing one or more adsorbent layers in a combined TSA/PSA process. Afirst layer of activated alumina or zeolite is commonly used for waterremoval and a second layer of zeolite such as 13X molecular sieve isused for CO₂ removal. Prior art, such as U.S. Pat. No. 4,711,645,teaches the use of various adsorbents and methods for removal of CO₂ andwater vapor from air. These adsorbents are ineffective for the removalof CO and H₂, thus allowing CO and H₂ to pass through to thedistillation equipment.

There are three principal strategies in the prior art to remove COand/or H₂ from air to produce UHP nitrogen: removal upstream of theprepurifier adsorber, removal within the prepurifier adsorber using anoxidation catalyst and removal from the nitrogen product after cryogenicair separation.

In the first approach, CO and H₂ are usually removed by high temperaturecatalytic oxidation over a supported noble metal or hopcalite catalystupstream of the prepurifier beds. The products from oxidizing CO and H₂,namely CO₂ and H₂O, are removed along with the ambient CO₂ and H₂O inthe prepurifier beds (F. C. Venet, et al., “Understand the Key Issuesfor High Purity Nitrogen Production,” Chem. Eng. Prog., pp 78-85,January 1993). This approach requires significant power and additionalcapital, adding substantially to the cost of the process.

U.S. Pat. No. 5,656,557 discloses a process wherein the compressed airis further heated to 350° C. prior to entering a catalyst towercontaining palladium (Pd) and/or platinum (Pt) supported catalyst forconverting CO, H₂ and hydrocarbons to H₂O and CO₂. The processed air isthen cooled to 5° C. to 10° C. prior to entering the prepurifier wherethe H₂O and CO₂ are removed. Part of the effluent from the prepurifiermay be used as air containing less than 1 ppm total impurities, whilethe remaining air is separated cryogenically to produce N₂ and O₂.

French patent FR 2 739 304 describes a method of removing CO and H₂ fromair which involves; 1) contacting the compressed hot moist gas from thecompressor with a bed of CO oxidation catalyst; 2) cooling the resultingintermediate air stream to ambient temperature; 3) contacting this COfree stream with an adsorbent to adsorb CO₂ and H₂O; and 4) contactingthe resulting stream with a H₂ trapping adsorbent. The CO catalyst canbe copper (Cu) or a Pt group metal supported on alumina, silica orzeolite. The H₂ trapping adsorbent can be osmium (Os), iridium (Ir), Pd,ruthenium (Ru), rhodium (Rh) or Pt supported on alumina, silica orzeolite.

U.S. Pat. No. 6,074,621 describes a similar process as FR 2 739 304except for the cooling step after the CO oxidation catalyst.

U.S. Pat. No. 5,693,302 discloses a method of removing CO and H₂ from acomposite gas by passing over particles containing gold and Pd supportedby TiO₂.

U.S. Pat. No. 5,662,873 describes a similar process using a catalystconsisting of silver and at least one element from Pt family supportedon alumina, silica or zeolite.

A second technology employed in the prior art is an ambient temperatureprocess for CO and H₂ removal from air.

U.S. Pat. No. 5,110,569 discloses a process for removing CO andoptionally hydrogen from air by 1) removing water 2) catalyticallyoxidizing CO to CO₂ and optionally H₂ to H₂O and 3) removing theoxidation products. Oxidation catalysts for CO can be a mixture ofmanganese and copper oxides such as hopcalite or Carulite. Nickel oxideis also stated to be an effective CO catalyst. The oxidation catalystfor H₂ is typically supported palladium.

U.S. Pat. No. 5,238,670 discloses a method of removing CO and/or H₂ fromair at a temperature between 0° C. and 50° C. by 1) removing water fromair until it has a water content lower than 150 ppm and 2) removing COand H₂ on a bed of particles containing at least one metallic elementselected from Cu, Ru, Rh, Pd, Os, Ir and Pt deposited by ion-exchange orimpregnation on zeolite, alumina or silica.

European patent application EP 0 454 531 describes a similar methodwhich suggests removing both H₂O and CO₂ prior to the impregnated bed ofparticles. Traces of H₂O and CO₂ are removed downstream of theimpregnated particle bed.

U.S. Pat. No. 6,048,509 discloses a method for removing CO and H₂ fromair at ambient temperature wherein air containing H₂O, CO₂, CO andoptionally H₂ passes through following steps; 1) contacting the gas witha CO catalyst consisting of Pd or Pt and at least one member selectedfrom the group consisting of iron (Fe), cobalt (Co), nickel (Ni),manganese (Mn), Cu, chromium (Cr), tin (Sn), lead (Pb), and cerium (Ce)supported on large pore alumina; 2) contacting the CO free gas with anadsorbent for water removal; 3) contacting the resulting gas with a CO₂adsorbent for CO₂ removal and optionally; 4) contacting the gas with aH₂ catalyst which consists of Pt or Pd supported on activated alumina orzeolite. The water formed in the last step of hydrogen oxidation iseither adsorbed on the H₂ catalyst support or it is removed by a H₂Oadsorbent, which is either physically mixed with the H₂ catalyst orplaced downstream of it.

U.S. Pat. No. 6,093,379 describes a method where a prepurifier bed witha first layer of water adsorbent and a second layer of CO₂ adsorbentoperating at ambient temperature is augmented by a third layer ofcatalyst/adsorbent to remove both CO and H₂. The third layer is exposedto substantially H₂O-free and CO₂-free air at ambient temperature. Thedual catalyst/adsorbent is placed in the third layer in the mostdownstream end of the prepurifier beds. The dual catalyst/adsorbentlayer oxidizes CO, adsorbs the resulting CO₂, and chemisorbs H₂. Thisdual catalyst/adsorbent is a precious metal such as Pd on a supporthaving a zero point charge (ZPC) of greater than 8.

U.S. Pat. No. 6,511,640 discloses a method wherein a prepurifier isconfigured to contain various materials layered in series beginning withan adsorbent at the feed inlet for H₂O removal. The second layer, anoxidation catalyst to convert CO to CO₂, is followed by an adsorbent forCO₂ removal. An oxidation catalyst is placed in the next layer toconvert H₂ to H₂O, while the final layer is used for adsorbing H₂O. TheCO catalyst disclosed is hopcalite, while the H₂ catalyst is Pdsupported on activated alumina.

A third common strategy for producing UHP N₂ in the prior art is thetreatment of the cryogenically separated N₂ product to remove H₂, CO, O₂and other contaminants penetrating the prepurifier and air separationunit.

U.S. Pat. No. 4,579,723 discloses the use of a Ni or Cu supportedcatalyst or getter to oxidize the contaminants to CO₂ and H₂O, which aresubsequently removed in an adsorber.

European patent EP 0 835 687 teaches regeneration of catalyst beds witha high temperature N₂ purge.

Adsorption of CO has been applied in the prior art predominantly forrecovery of CO in bulk separations, e.g. in cases where theconcentration (partial pressure) of CO is relatively high (typically=1%)and where CO is the more strongly adsorbed component in the gas mixture.Cuprous compounds, either in cationic form in zeolites or dispersed on aporous support, have been widely applied in the recovery of CO from gasmixtures containing CO and N₂, methane (CH₄), H₂ and/or CO₂. Materialscontaining copper in the single oxidation state (denoted as Cu⁺, Cu(I)or cuprous) display high CO adsorption capacity, while adsorbentscontaining Cu(II) do not. Adsorbents are commonly synthesized, treatedor modified with a Cu(II) compound and then subsequently exposed to areducing agent such as H₂ at elevated temperature to convert the Cu(II)to Cu(I).

Xie et al. (“Highly Efficient Adsorbent for Separation of CarbonMonoxide,” Fundamentals of Adsorption, Proc. IV^(th) Int. Conf. OnFundamentals of Adsorption, Kyoto, May 17-22, 1992, pp. 737-741)describes an adsorbent formed by dispersing CuCl on a zeolite support bymixing the dry powders at elevated temperature. High purity CO separatedto high recovery is demonstrated for feed streams containing 9.0% CO/91%N₂ and 30.7% CO/65.3% H₂/4% CH₄.

U.S. Pat. No. 4,917,711 discloses adsorbents and processes utilizingsupported CuCl. U.S. Pat. No. 5,531,809 discloses VSA processes usingCuCl dispersed on alumina for recovery of CO from synthesis gas exitinga steam-methane reformer.

G. K. Pearce (“The Industrial Practice of Adsorption,” in: Separation ofGases, 5^(th) BOC Priestley Conf., Birmingham, UK Sep. 19-21, 1989,Spec. Publ. No. 80, Royal Soc. Of Chemistry, Cambridge, 1990) provides adescription on the use of Cu(I)Y zeolite for the recovery of CO fromCO/N₂ and CO/H₂ feed streams containing percentage (%) levels of CO.

U.S. Pat. No. 4,473,276 discloses Cu(I)Y and Cu-Mordenite along withother exchanged zeolites having a silica to alumina ratio(SiO₂/Al₂O₃)=10 for the recovery of CO.

U.S. Pat. No. 4,019,879 discloses recovery of CO from streams containingH₂O and/or CO₂ using Cu⁺ containing zeolites with 20=SiO₂/Al₂O₃₌₂₀₀,e.g. ZSM-5, -8, -11, etc.

Another class of adsorbents having potential for CO adsorption is one inwhich the materials contain silver (Ag⁺). Y. Huang (“Adsorption in AgXand AgY Zeolites by Carbon Monoxide and Other Simple Molecules,” J.Catal., 32, pp. 482-491, 1974) provides CO and N₂ isotherms for AgX andAgY zeolites to partial pressures only as low as 0.1 to 1.0 torr for thelowest temperature (0° C. and 25° C.) isotherms. The adsorption capacityfor CO is significantly greater than that of N₂ at 25° C. and 100 torr.

U.S. Pat. No. 4,743,276 discloses mordenite, A, Y and X type zeolitesexchanged with various amounts of Ag for the bulk separation (recovery)of CO from refinery and petro-chemical off-gases.

U.S. Pat. No. 4,019,880 relates to the recovery of CO from gas streamscontaining also H₂O and/or CO₂ using Ag exchanged zeolites with20=SiO₂/Al₂O₃=200, e.g. ZSM-5, -8, -11, etc. The invention applies tofeed streams containing at least 10 ppm CO at temperatures 0° C.-300° C.The claimed process results in a CO-depleted effluent, e.g. air.

U.S. Pat. No. 4,944,273 discloses zeolites with 1=Si/Al=100 and dopedwith Ca, Co, Ni, Fe, Cu, Ag, Pt or Ru for adsorption of oxides ofnitrogen (NO_(x)) and CO as part of NO_(x) and CO sensors, particularlyin exhaust gases of automotive vehicles.

U.S. Pat. No. 3,789,106 discloses that mordenite charged with copper iseffective in removing CO from H₂ at CO partial pressure below 3 mmHg.The effectiveness was determined by subjecting the adsorbent to COconcentrations greater than or equal to 100 ppm and measuring capacityat saturation.

The above prior art relating to adsorption of CO is almost totallysilent with respect to purification of CO from mixed gas streams,particularly those containing less than 10 ppm CO in O₂ and N₂.

OBJECTS OF THE INVENTION

It is therefore an object of the invention to provide an improved methodfor the removal of trace amounts of CO and at least one of H₂, H₂O, CO₂,hydrocarbons and N₂O from feed gas streams, preferably air.

SUMMARY OF THE INVENTION

The invention relates to the removal of CO and optionally H₂ from airand/or other gases or gas mixtures using a combination of adsorptiveseparation and catalytic conversion, and provides unexpected savings incapital and power costs over existing technologies.

In one preferred embodiment, at least 90% of each of CO₂ and H₂O arefirst removed from the feed gas (preferably air) to produce a CO₂ andH₂O depleted gas. In a particularly preferred embodiment, the CO₂ andH₂O depleted gas contains less than 1.0 ppm (more preferably less than0.25 ppm) CO₂ and less than 1.0 ppm (more preferably less than 0.10 ppm)H₂O. A substantial amount of CO is then removed from the CO₂ and H₂Odepleted gas through the use of a CO adsorbent to produce a CO depletedgas containing, in a preferred embodiment, less than 100 ppb CO, morepreferably less than 5 ppb CO. Optionally H₂ and any remaining CO maythen be removed using a catalyst. Because of this unique combination ofcatalysis and adsorption, the process of the present invention providessurprisingly superior CO and H₂ removal efficiency over the prior artprocesses.

A preferred apparatus for the practice of the invention comprises anadsorption apparatus for the removal of CO from a feed stream containingCO in an amount of less than 50 ppm. The apparatus comprises at leastone adsorption vessel containing a CO adsorbent layer, the CO adsorbenthaving a ΔCO working capacity greater than or equal to 0.01 mmol/g; andwherein

a) when the feed stream further contains at least one gas selected fromthe group consisting of nitrogen, He, Ne, Ar, Xe, Kr, CH₄ and mixturesthereof, the adsorbent is ion exchanged with a Group IB element, and ispreferably selected from the group consisting of AgX zeolite,Ag-Mordenite, Cu-clinoptilolite, AgA zeolite and AgY zeolite, and

b) when the feed stream further contains at least one gas selected fromthe group consisting of oxygen and air and mixtures thereof, theapparatus is preferably an air prepurifier, and the adsorbent is azeolite having a SiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with aAg⁺ or Au⁺, and is preferably selected from the group consisting of AgXzeolite, Ag-Mordenite, AgA zeolite and AgY zeolite.

In a further embodiment the apparatus contains two or more adsorptionvessels that are selected from the group consisting of vertical flowvessels, horizontal flow vessels, lateral flow vessels or radial flowvessels.

In a further embodiment the adsorption apparatus further contain anadsorbent, preferably one or more of alumina or NaX zeolite, selectivefor the adsorption of water that is upstream of said CO adsorbent layer.

In a further embodiment the apparatus further contains a catalyst layer,preferably a metal supported catalyst comprising one or more of themetals Os, Ir, Pt, Ru, Rh, Pd, Fe, Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr,Pb, Ce supported on a substrate selected from the group consisting ofalumina, silica, natural or synthetic zeolites, titanium dioxide,magnesium oxide and calcium oxide, for the catalytic oxidation of H₂ toH₂O that is downstream of said CO adsorbent layer.

In a further embodiment the apparatus further contains an auxiliaryadsorbent, preferably one or more of alumina or NaX, for the removal ofwater that is downstream of said catalyst layer.

In a further embodiment the ΔCO working capacity is greater than orequal to 0.03 mmol/g.

In a further embodiment, the apparatus further contains one or moreadditional adsorbents for the adsorption of one or more of H₂O, CO₂, N₂Oand hydrocarbons, wherein the additional adsorbents preferably areselected from the group consisting of alumina, silica gel,clinoptilolite, zeolites, composites thereof and mixtures thereof, andare located downstream of the catalyst layer.

In a further embodiment the CO adsorbent has a ΔCO/ΔN₂ separation factorof greater than or equal to 1×10⁻³.

In a further embodiment the CO adsorbent is AgX having at least 50%,preferably 100% of its cations associated with Ag.

A preferred process for the practice of the invention comprises aprocess for the removal for the removal of CO from a feed streamcontaining CO in an amount of less than 50 ppm, or even less than 1.0ppm or 0.5 ppm. The process, which is preferably an air prepurificationprocess, comprises contacting the feed stream with a CO adsorbent havinga ΔCO working capacity greater than or equal to 0.01 mmol/g, preferablygreater than or equal to 0.03 mmol/g, to produce a CO depleted gasstream, that may be recovered; and wherein

a) when the feed stream further contains at least one gas selected fromthe group consisting of nitrogen, He, Ne, Ar, Xe, Kr, H₂, CH₄ andmixtures thereof, the adsorbent is a zeolite exchanged with a Group IBelement and is preferably selected from the group consisting of AgXzeolite, Ag-Mordenite, Cu-clinoptilolite, AgA zeolite and AgY zeolite,and

b) when the feed stream further contains at least one gas selected fromthe group consisting of oxygen and air and mixtures thereof, theadsorbent is a zeolite having a SiO₂/Al₂O₃ ratio of <20, and ision-exchanged with a Ag⁺ or Au⁺, and is preferably selected from thegroup consisting of AgX zeolite, Ag-Mordenite, AgA zeolite and AgYzeolite.

In a further embodiment, the process comprises recovering the COdepleted gas stream, wherein CO is present in the CO depleted gas streamat a concentration of less than 100 ppb, preferably less than 5 ppb,most preferably less than 1 ppb.

In a further embodiment, said feed gas further comprises water (H₂O),and the process further comprises contacting the feed stream with awater selective adsorbent, preferably alumina or NaX, that is locatedupstream of said CO adsorbent.

In a further embodiment, the feed gas further comprises hydrogen, andthe process further comprises contacting the CO depleted feed streamwith a catalyst layer that is preferably a metal supported catalystcomprising one or more of the metals Os, Ir, Pt, Ru, Rh, Pd, Fe, Co, Ni,Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce supported on a substrate selectedfrom the group consisting of alumina, silica, natural or syntheticzeolites, titanium dioxide, magnesium oxide and calcium oxide, for thecatalytic oxidation of H₂ to H₂O to produce a H₂ depleted and H₂Oenriched gas, and wherein the catalyst layer is located downstream ofsaid CO adsorbent layer. The hydrogen depleted gas preferably containsless than 100 ppb hydrogen, and more preferably less than 5 ppbhydrogen.

In another embodiment, the process further comprises the step ofcontacting the catalytically produced H₂O enriched gas with an adsorbentfor the removal of water (preferably alumina or NaX) that is locateddownstream of said catalyst layer to produce a gas that is depleted inCO, H₂ and H₂O.

In another embodiment the CO adsorbent has aΔCO/AN₂ separation factorthat is greater than or equal to 1×10⁻³, preferably greater than orequal to 1×10⁻².

In another embodiment, the process further comprises passing said feedgas over

a) at least one adsorbent layer upstream of said CO adsorbent for theadsorption of one or more of H₂O and CO₂,

b) a catalyst layer for the catalytic conversion of H₂ to H₂O that isdownstream of said CO adsorbent layer; and

c) one or more additional adsorbents, preferably selected from the groupconsisting of alumina, silica gel, clinoptilolite, zeolites, compositesthereof and mixtures thereof, for the adsorption of one or more of H₂O,CO₂, N₂O and hydrocarbons, wherein said additional adsorbents aredownstream of said catalyst layer.

In an alternative embodiment, the process is selected from the groupconsisting of pressure swing adsorption, temperature swing adsorption,or a combination thereof, and the process takes place in an adsorbervessel selected from a vertical flow vessel, a horizontal flow vessel ora radial flow vessel.

In a preferred embodiment the adsorption step of the process is operatedat a temperature of zero to fifty degrees Celsius.

In a preferred process the CO adsorbent is AgX having at least 50%, morepreferably 100% of its cations associated with Ag.

In a preferred embodiment the feed stream contains air, and the COdepleted gas stream is passed to a cryogenic distillation column.

In an alternative embodiment the CO partial pressure in the feed streamis less than 0.1 mmHg, or even less than 0.005 mmHg.

An alternative embodiment relates to a process for the removal for theremoval of CO from a feed stream containing CO in an amount of less than50 ppm and hydrogen, the process comprising contacting the feed streamwith a CO adsorbent having a ΔCO working capacity greater than or equalto 0.01 mmol/g to produce a CO depleted gas stream; and wherein theadsorbent is a zeolite exchanged with a Group IB element.

In a preferred embodiment, the invention comprises an adsorptionapparatus for the removal of CO from a feed stream containing CO in anamount of less than 50 ppm, and hydrogen, the apparatus comprising atleast one adsorption vessel containing a CO adsorbent layer having a ΔCOworking capacity greater than or equal to 0.01 mmol/g; and wherein saidadsorbent is a zeolite exchanged with a Group IB element.

In a particularly preferred embodiment, the invention comprises an airprepurification adsorption apparatus for the removal of CO from an airfeed stream containing CO in an amount of less than 50 ppm, theapparatus comprising at least one adsorption vessel containing a COadsorbent layer having a ΔCO working capacity greater than or equal to0.01 mmol/g; and wherein the adsorbent is a zeolite (preferably AgX),having a SiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with a Ag⁺ orAu^(+.)

In a preferred embodiment, the air prepurification apparatus furthercomprises at least one adsorbent layer upstream of said CO adsorbent forthe adsorption of one or more of H₂O and CO₂.

In a preferred embodiment, the prepurification apparatus furthercomprises a catalyst layer for the catalytic conversion of H₂ to H₂Othat is downstream of the CO adsorbent layer; and one or more additionaladsorbents for the adsorption of one or more of H₂O, CO₂, N₂O andhydrocarbons, wherein said additional adsorbents are downstream of saidcatalyst layer.

In an alternative embodiment, the invention comprises a process for theremoval of CO from a feed stream containing CO in an amount of less than50 ppm and air said process comprising contacting said feed stream in anadsorber vessel with a CO adsorbent having a ΔCO working capacitygreater than or equal to 0.01 mmol/g to produce a CO depleted gasstream; and wherein the adsorbent is a zeolite having a SiO₂/Al₂O₃ ratioof <20, and is ion-exchanged with a Ag⁺ or Au⁺, preferably AgX. Thisprocess may further comprise a) passing the feed stream over at leastone adsorbent layer upstream of said CO adsorbent for the adsorption ofone or more of H₂O and CO₂ to produce a H₂O and or CO₂ depleted feedstream b) passing the feed stream over a catalyst layer that isdownstream of said CO adsorbent layer for the catalytic conversion of H₂to H₂O layer and one or more additional adsorbents for the adsorption ofone or more of H₂O, CO₂, N₂O and hydrocarbons, wherein said additionaladsorbents are downstream of said catalyst layer to produce a feedstream that is depleted in CO, H₂ and one or more of H₂O, CO₂, N₂O andhydrocarbons, and, optionally, passing the depleted feed stream to acryogenic distillation column for the separation of air.

Combinations of any of the above embodiments are contemplated to bewithin the scope of the invention.

BRIEF DESCRIPTION OF THE DRAWING

Other objects, features and advantages will occur to those skilled inthe art from the following description of preferred embodiment(s) andthe accompanying drawings, in which:

FIG. 1 is a schematic diagram of a breakthrough test apparatus;

FIG. 2 is a graph of CO and H₂ breakthrough curves for the AgX adsorbentof Table II;

FIGS. 3-6 are schematic diagrams of preferred adsorbent arrangements inan adsorbent vessel/bed.

FIG. 7 is a schematic diagram of a prepurification apparatus useful forcarrying out the invention.

DETAILED DESCRIPTION OF THE INVENTION

The invention relates to the removal of CO and optionally H₂ from a feedgas containing ppm levels of CO and ppm levels of H₂. The invention isparticularly applicable to removing CO from feed gas streams having 1ppm or less CO therein; in particular, feed gas streams having COpartial pressures of less than 0.1 mmHg or even as low as 0.005 mmHg orlower.

The feed gas is typically air, and may also contain other contaminantsincluding at least one of CO₂, H₂O, N₂O and hydrocarbons such asacetylene. The removal of these components typically takes place in aprepurification process prior to air separation by cryogenic means,though the invention may also be applied to remove CO and, optionallyhydrogen impurities after cryogenic distillation of air.

The invention may be practiced with feed gas streams containing as muchas 50 ppm CO and 10 ppm H₂ in the feed, though feed gas streamscontaining less than 1 ppm or even less than 0.5 ppm CO and 1 ppm oreven less than 0.5 ppm H₂ may also be used. Practice of the inventionwith such feed streams can produce a product gas stream containing lessthan 5 ppb CO and less than 5 ppb H₂, though gas streams containing 1ppb CO and 1 ppb H₂ can also be achieved.

The use of a CO adsorbent followed by a H₂ catalyst in accordance with apreferred embodiment of the invention reduces the H₂ catalystrequirement, eliminates the need for high temperature activation,maximally protects the H₂ catalyst from poisoning by locating it at theclean end of the adsorber and reduces capital and operation costs ascompared to existing prepurification techniques. The result is that theeffluent from the inventive adsorber/adsorption process (e.g. the feedto a cryogenic distillation column (or “cold box”) of an air separationunit (ASU)) contains less than 5 parts per billion (ppb) CO and lessthan 5 ppb H₂.

In one preferred embodiment of the invention, an adsorbent bed isconfigured to first remove from a feed stream containing CO, H₂, CO₂ andH₂O impurities, substantially all of the H₂O and optionally all of theCO₂ prior to removing CO on a CO adsorbent. The CO₂, H₂O and CO depletedgas is then passed through a catalyst bed to remove any remaining CO andthe H₂. The thus purified gas is then passed to an ASU to produce UHPnitrogen. For the purpose of this disclosure, a “high purity” gas meansa gas having individual contaminant levels of less than about 100 ppb(parts per billion) and an “ultra high purity” (UHP) gas means a gashaving individual contaminant levels less than about 10 ppb, preferablyless than about 5 ppb and most preferably less than 1 ppb, dependingupon the intended application of the gas.

Thus the present invention produces a gas depleted of H₂ and CO in a twostep process:

1) Removal of CO by adsorption from a feed stream containing traceamounts of CO at temperatures in the range of 0 to 50 degrees Celsius(C.) on a suitable adsorbent in the presence of high N₂ and/or O₂concentrations to produce an effluent gas containing less than 5 ppb CO;and

2) optional removal of trace amounts of H₂ from a feed stream in thepresence of high N₂ and/or O₂ concentrations by a combination ofadsorption and/or absorption and/or catalysis at temperatures in therange of 0 to 50 degrees C. to produce an effluent gas containing lessthan 5 ppb H₂.

In accordance with one embodiment of the invention, an appropriate COadsorbent is selected that has sufficient ΔCO working capacity to removemost or all of the CO from the feed stream. An ideal CO adsorbent wouldhave good CO loading (ΔCO or working capacity) in the presence of highN₂ and/or O₂ concentrations. It is also desirable to have low N₂ loadingto maximize CO selectivity and to minimize the cooling effect of N₂desorption during depressurization. In a preferred embodiment theadsorbent also has a high ΔCO/AN₂ working separation factor at the localstream conditions.

One method for estimating adsorbent performance is by determining theworking capacity of each of the primary adsorbates, i.e. N₂ and CO. Thisis the methodology applied in the present invention. The separationfactor α, as defined below, is a preferred way to evaluate the adsorbenteffectiveness. This methodology is discussed in detail in U.S. Pat. No.6,152,991, the contents of which are herein incorporated by reference.

Separation factor α is defined as follows: $\begin{matrix}{\alpha = {\frac{\Delta\quad{CO}}{\Delta\quad N_{2}} = \frac{{w_{CO}\left( {y,p,T} \right)}_{ads} - {w_{CO}\left( {y,p,T} \right)}_{des}}{{w_{N_{2}}\left( {y,p,T} \right)}_{ads} - {w_{N_{2}}\left( {y,p,T} \right)}_{des}}}} & (1)\end{matrix}$

where separation factor α is defined as the ratio of the workingcapacities. The numerator in this equation is the working capacity ofCO, which is equal to the difference in loading w between adsorption anddesorption conditions. The adsorption and desorption conditions arecharacterized by composition y, pressure p and temperature T. In TSA airprepurification, maximum regeneration temperatures may vary from about100° C. to about 350° C. As a result, it is expected that the adsorbates(particularly atmospheric gases) will be completely thermally desorbedfrom the adsorbents. Under such conditions, Equation (1) can besimplified as follows: $\begin{matrix}{\alpha = {\frac{\Delta\quad{CO}}{\Delta\quad N_{2}} = \frac{{w_{CO}\left( {y,p,T} \right)}_{ads}}{{w_{N_{2}}\left( {y,p,T} \right)}_{ads}}}} & (2)\end{matrix}$

When a contaminant is removed in a shallow adsorbent layer in a TSAprocess and significant resistance to mass transfer exists, theselectivity is redefined according to Equation (3): $\begin{matrix}{\frac{\Delta\quad X_{CO}}{\Delta\quad X_{N_{2}}} = \frac{\frac{m_{in}}{w_{s}}{\int_{0}^{t_{b}}{\left( {y_{in} - y_{out}} \right){\mathbb{d}t}}}}{{X_{N_{2}}\left( {y,P,T} \right)}_{ADS}}} & (3)\end{matrix}$The numerator in Equation (3) represents the working capacity of theadsorbent for the contaminant. m_(in) represents the molar feed flowinto the bed, y_(in) and y_(out) are the inlet and outlet mole fractionsof the minor component, respectively. w_(s) is the mass of adsorbent andt_(b) is the breakthrough time corresponding to a predeterminedconcentration. The denominator is the equilibrium capacity of the majorcomponent at the conditions at the end of the adsorption step, i.e.assuming complete desorption of all components. Application of Equation(3) is preferred when the length of the mass transfer zone is anappreciable fraction (e.g. more than about 10%) of the overall depth ofthe adsorbent layer.

This method is preferred over other methods because the workingcapacities are determined at the partial pressure of each individualcomponent at the relevant process conditions, e.g. in the case ofprepurification the CO and N₂ partial pressures in the feed air aretypically <0.1 mmHg and >1000 mmHg, respectively. Coadsorption effectsare also incorporated into the determination of the loading. Since inair prepurification the feed concentration of N₂O is overwhelmingcompared to CO, the coadsorption effect of CO upon N₂ is negligible.Thus, the denominator of Equation (2) and Equation (3) may be obtaineddirectly from the measured pure-component N₂ isotherm.

Conversely, the coadsorption of N₂ would have a very significant effectupon the adsorption of CO. If accurate low concentration pure-componentisotherm data for CO is available or attainable, then Equation (2) maybe applied to assess equilibrium working capacity and selectivity.However, it is preferred to determine the working capacity for COdirectly under N₂ coadsorption conditions and with mass transfer effectsincluded. This may be done using a breakthrough test method, which iswell known to those skilled in the art. Breakthrough tests provide theequilibrium capacity of a component at saturation, and the breakthroughcapacity and time at some defined breakthrough level, e.g. 100 ppb.Thus, the ΔCO working capacity and selectivity are preferably determinedusing the terms of Equation (3). Since the coadsorption of nitrogen isincorporated into the ΔCO working capacity, ΔCO determined by thebreakthrough method (numerator of Equation (3)) is the dominant factorin selecting an appropriate Co adsorbent for the purpose of theinvention.

One skilled in the art will appreciate that adsorbents satisfying thecriteria for acceptable working characteristics for CO relative to N₂will perform even more favorably in other gas mixtures where the bulkcomponent adsorbs less strongly than N₂. Examples of such CO-containinggas mixtures include those with one or more of He, H₂, Ar, Ne, Xe, Kr,O₂ and CH₄, and the invention may also be applied to the separation ofCO from these gases or gas mixtures.

In order to evaluate the potential effectiveness of CO adsorbents atconditions typical of air separation processes, CO and/or H₂breakthrough tests were performed.

Breakthrough tests were performed at 7.9 bara (114.7 psia), either 10°C. or 27° C. and an inlet gas flowrate of approximately 21 slpm (78.7mol/m² s) using an adsorption column length of 5.9 cm or 22.9 cm.Variations to these conditions are noted in the examples. The feedconditions are representative of conditions at the inlet of an airprepurifier for a typical cryogenic air separation plant. Completebreakthrough curves were generated for CO and/or H₂ in some of thetests, while only partial breakthroughs were determined in those casesusing high capacity materials. Initial breakthrough was established at100 ppb CO and 20 ppb H₂. Initial breakthrough and/or saturation loading(capacity in mmol/g) were then determined according to the mass balancesas indicated in the numerator of Equation (3). CO and H₂ loadingdetermined at the initial breakthrough represent a dynamic or workingcapacity—incorporating both coadsorption and mass transfer effects.

Breakthrough tests were conducted in the following manner using anapparatus for which the key elements are shown in the schematic ofFIG. 1. Challenge gases from cylinders (3,4) containing a mixture of thecontaminant gas in N₂ were metered through flow controllers (7,8) (atflow rates on the order of 0-5 standard liters per minute (slpm)) wherethey were mixed in a gas mixer (9) with high purity diluents N₂ orhelium (He). In some cases the diluent was synthetic air wherein O₂ fromsource 2 was combined with N₂ from source 1. This diluent was providedat a prescribed flowrate (e.g. 0-30 slpm) through flow controllers (5,6)to achieve the desired feed concentration of contaminant(s). This mixedchallenge gas was then fed through a heat exchange loop and to the testbed 10 containing the adsorbent. Both the test bed 10 and heat exchangeloop were located within a thermal bath (not shown), the temperature ofwhich was controlled via a water chiller and thermocouples. Temperaturecontrol systems are well known in the art and typically consist of aheat exchange loop, thermal bath, water chiller and thermocouples. Anynumber of variations of the temperature control system can beeffectively applied to maintain the test bed at a constant temperatureas would be readily familiar to one of ordinary skill in the art.

The gas pressure during the test was controlled through a controlvalve/pressure controller 11. A portion of the effluent was passedthrough valve 12 to one or more analyzers 13 to monitor the breakthroughconcentrations of CO and/or H₂ as a function of time. A Servomex 4100Gas Purity Analyzer equipped with both CO and CO₂ sensors was used inExample 1. In Examples 2-4 a Trace Analytical (RGA5) Analyzer was usedfor trace CO and H₂ measurement.

All zeolites were obtained from commercial sources (identified below),but some were further modified by the inventors by ion exchange.Mordenite and Type Y zeolites were extrudates, Type X and Type Azeolites were beads and clinoptilolite and chabazite were granular forthe examples of this invention.

A laboratory ion-exchange column was used to create highly exchangedzeolites. In the column procedure, the zeolite was packed inside thecolumn and ion exchange solution was pumped upwards through the bed at aflow rate of 0.5 ml/minute. The column was heated between 70-90° C. topromote exchange. At least a ten-fold excess of solution was used toensure that high levels of ion exchange are obtained. The product wascollected by filtration and washed thoroughly with deionized water.

Powder X-ray diffraction and inductively coupled plasma (ICP) chemicalanalysis were used to verify the integrity of the samples as well asdetermine the level of ion exchange. For the Ag, Cu and zinc (Zn)samples, a 0.1 molar (M) solution of the respective nitrate salt wasused. For sodium (Na) exchange, a 1.0M sodium chloride salt solution wasused. The zeolites were typically in their Na exchanged forms prior toexchange with Ag, Cu or Zn. For clinoptilolite, a TSM-140 sample wasselected and exchanged first with Na before Ag, Cu or Zn was introduced.Zeolites 13X HP, 13X APG and Na mordenite were obtained from UOP andzeolite 4A was purchased from Aldrich.

For lower exchange levels of Ag on 13X HP (e.g. <85%), a batch procedurewas used. In the batch process the zeolite was immersed in a fixedvolume of solution as opposed to having fresh solution pumped into acolumn on a regular basis. Silver nitrate solution having aconcentration of 0.025-0.1M was used.

The zeolites were stirred in a solution of silver nitrate at 50-90° C.The solution strength was chosen to reflect the level of exchangedesired. After a period of 6-8 hours stirring at temperature, thesolution was either refreshed up to a total of three times, or thesample was collected by filtration and washed thoroughly with deionisedwater. Refreshing the solution increased the exchange level. PowderX-ray diffraction and ICP chemical analyses were performed to examinethe crystallinity and ion exchange level of the product samples.

Methods for ion exchange of zeolites are well known to those of ordinaryskill in the art. The column and batch methods described above are in noway limiting to the invention. Different procedures may be used in orderto achieve desired exchange levels for the cations and zeolites inaccordance with the present invention.

The invention will now be described with reference to the followingcomparative examples. These examples are not intended to limit the scopeof the invention in any way.

EXAMPLE 1 Outside the Scope of the Invention

The natural adsorbents (clinoptilolite and chabazite) were obtained fromSteelhead Specialty Minerals, WA. Synthetic zeolites were obtained fromvarious manufacturers: Zeolyst (HZSM5), Zeochem (CaX) and UOP (13X, LiX,5AMG). All of the adsorbents were thermally activated at 350° C., 1.0bara pressure and under N₂ purge for approximately 16 hours before eachtest. After regeneration the adsorbents were allowed to cool to 27° C.This group of adsorbents presents a cross section of adsorbentcharacteristics known to effect adsorption properties, e.g. Si/Al,micropore channel opening size, zeolite structure type, number and typeof cations, etc. Breakthrough tests were performed at the conditionsdescribed above to saturation using 1.0 ppm CO in N₂ or He. Breakthroughtime was determined at 100 ppb CO. The results are summarized in TableI. TABLE I Flow T y_(co) Bed t_(b) ⁽²⁾ X_(co) ⁽³⁾ Adsorbent slpm ?CCarrier⁽¹⁾ ppm cm min mmol/g Clinoptilolite 20.4 27 N₂ 1.0 22.9 <2 1.4 ×10⁻⁵ TSM-140 21 27 He 1.0 22.9 30 1.2 × 10⁻³ Chabazite 21 27 N₂ 1.0 22.9<2 2.2 × 10⁻⁵ TSM-300 21 27 He 1.0 22.9 114  4.5 × 10⁻³ LiX 21 27 N₂ 1.022.9 <2 2.5 × 10⁻⁵ (SiO₂/ Al₂O₃ = 2.0) 5A MG 5.5 10 N₂ 1.0 22.9 <6 3.1 ×10⁻⁵ CaX (Z100) 5.5 10 N₂ 1.0 22.9 <4 3.9 × 10⁻⁵ HZSM5 21 27 air 2.0 5.9 <2 — (CBV3024E)⁽¹⁾typical H₂ concentrations in carrier gases .0.25 ppm⁽²⁾breakthrough determined at y_(co) = 100 ppb⁽³⁾equilibrium saturation capacity of CO

The competitive effect of N₂ upon CO saturation capacity (X_(CO)) isevident in that X_(CO) for CO/N₂ feed was reduced by a factor of 100 ormore compared to that for CO/He feed for the small pore naturalzeolites. Breakthrough times of CO in N₂ were reduced by factors of 15to more than 50 compared to CO in He. The high partial pressure of N₂relative to that of CO is a critical factor in determining adsorbenteffectiveness in CO removal. All of the adsorbents in Table I have CO/N₂breakthrough times of only a few minutes. Although breakthrough time canbe extended with longer adsorbent beds, none of these adsorbents islikely to provide a practical solution for air prepurifiers withadsorption cycle times of one hour or more.

N₂ isotherms were measured at 27° C. using well known gravimetricbalance methods for chabazite and clinoptilolite. The N₂ capacity fromthe isotherms and the CO saturation capacity from Table I were combinedin Equation 2 to result in equilibrium separation factors for theadsorbents in Table I (ΔCO/ΔN₂)<2×10⁻⁵.

EXAMPLE 2

Several zeolites containing exchanged Ag, Cu and Zn cations were testedas described above to evaluate the working CO capacity. AgX (P/N38,228-0) was obtained from Aldrich, while clinoptilolite (TSM-140) andmordenite (large-pore from UOP) were exchanged in small-scale laboratorycolumns. The extent of exchange was determined by inductively coupledplasma (ICP) analysis. Exchanged samples were air-dried and thenactivated overnight in a dry N₂ purge at 350° C. Breakthrough tests wereconducted using a 5.9 cm long column filled with adsorbent and subjectedto 2.0 ppm CO in synthetic air (79% N₂/21% O₂) at 27° C. and 7.9 bar,flowing at a rate of approximately 21 slpm (78.7 mol/m²s).

The amount of H₂ in the feed for each test is given in Table II alongwith the results of the breakthrough tests. The breakthrough time(t_(b)) and the working capacity (X_(CO)) were determined at a CObreakthrough concentration of 100 ppb. TABLE II X_(co) % CO Size @y_(co)= removal % US Y_(H2) 100 ppb t_(b) @y_(co) = Adsorbent exch. mesh ppmmmol/g hr 100 ppb AgX 100  10 × 14 3.0 0.052 5.4 98.8 Ag-Mor 89 1.8 mm*3.0 0.034 3.2 98.3 Cu- 100  8 × 14 0.5 0.037 3.0 99.3 clinoptilolite Zn-86 8 × 14 0.6 — <0.2  — clinoptilolite Ag- 89 8 × 14 0.4 — <0.5  —clinoptilolite*extrudate diameter

AgX, Ag-mordenite and Cu-clinoptilolite all display CO workingcapacities in excess of 0.03 mmol/g, and consequently meet the criteriaof the invention. Zn-clinoptilolite and Ag-clinoptilolite had workingcapacities less than 0.01 mmol/g and consequently are outside the scopeof the invention.

Conservatively compared to the CO saturation capacities in Table I, theCO working capacities in Table II (including dynamic effects) for theseadsorbents are more than 1000 times greater than those in Table I.Furthermore, breakthrough times are several hours for relatively shortbeds, i.e. the resultant CO working capacity allows the adsorbent to beeasily integrated with current prepurifier cycles and with a minimum ofadditional adsorbent.

Cu-, Zn- and Ag-exchanged clinoptilolite show little or no removalcapacity for H₂. Although H₂ breaks through almost immediately in AgXand Ag-mordenite, some modest holdup of H₂ is evident as can be seenfrom FIG. 2.

FIG. 2 gives the breakthrough history for both CO and H₂ using AgX(Aldrich P/N 38,228-0). Although a CO breakthrough concentration of 100ppb was chosen for the purpose of adsorbent evaluation, one skilled inthe art will appreciate that the amount of CO impurity in the productcan be adjusted by selecting either a shorter adsorption step or byincreasing the amount of adsorbent.

Beds were regenerated after each breakthrough test in air at 200° C. for3.0 hr followed by a ambient temperature purge in dry N₂ for 3.0 hr, allat a flow of 2.2 slpm and a pressure of about 1.7 bar (25 psia). Theresults in FIG. 2 are for the tenth breakthrough of this bed. There wasno noticeable deterioration in CO removal capacity with cycling—even inthe presence of the strong H₂ reducing agent. Desorption was monitoredafter the sixth breakthrough by purging the bed with N₂ at 27° C. and50° C. prior to the standard regeneration. Only CO was detected in theeffluent, and more than 87% of the adsorbed CO was desorbed. Theremaining CO was desorbed in the normal regeneration procedure.

N₂ isotherms were measured for AgX (Aldrich P/N 38,228-0) at 0° C. and27° C. as described above. The N₂ capacity at 27° C. at a N₂ partialpressure of 6.25 bar was determined as 0.81 mmol/g. The corresponding Copartial pressure for the test stream of Table II is 1.2×10⁻² mmHg. UsingEquation 3 and the results in Table II, the working separation factorΔCO/ΔN₂=6.4×10⁻².

The results in Table II for Cu-clinoptilolite are for the firstbreakthrough after activation. After regeneration following theconditions cited above, the performance was significantly degraded inthe second breakthrough. The performance in Table II was restored afterre-activating in N₂ at 350° C. It is believed that regeneration in airat elevated temperature resulted in the oxidation of Cu⁺ to Cu⁺⁺.Reactivation in N₂ returned the Cu to the lower oxidation state Cu⁺.Thus, this adsorbent may be more effectively applied in an alternativeembodiment of the invention which applies the inventive concept topost-purification of cryogenically separated N₂ where little or no O₂ ispresent.

EXAMPLE 3

Samples of commercially available AgX (Ag400B3), Ag-mordenite (Ag900E16)and AgY (Ag600E16) were obtained from C*CHEM, A Division of MolecularProducts, Inc., Lafayette, Colo. AgX was also prepared by exchanging 13XHP zeolite with Ag to various Ag-exchange levels (designated AgX (HP)).AgA was prepared by exchanging 4A zeolite with Ag. These materials andthe corresponding Ag exchange levels are given in Table III along withthe breakthrough time and CO working capacity. TABLE III Ag- X_(co)exch. Size @y_(co) = 100 ppb t_(b) % CO Adsorbent % US Mesh mmol/g hrremoval AgX 100  10 × 18 0.051 5.5 99.0 (Ag400B3) AgX (HP) 95 10 × 180.13 13.3 99.1 AgA 94  8 × 12 0.019 2.2 98.4 Ag-Mor 40 2.0 mm* 0.00960.85 97.9 AgY 61 1.7 mm* 0.018 1.4 98.8*extrudate diameter

Breakthrough tests were performed at the same conditions as those ofExample 2, except that 3.0 ppm H₂ was present in the feed for all tests.The results of Table III are for the first breakthrough afteractivation. The AgX (Ag400B3) performance agrees well with that of AgXin Table II. Although the AgX (HP) appears to have a significantlyimproved CO working capacity over the commercial AgX, much of thisadvantage was lost after several breakthrough/regeneration cycles. Afterthe fourth cycle, the breakthrough time for AgX (HP) was reduced to 7.3hr and the CO working capacity fell to 0.073 mmol/g. The amount ofdegradation, however, was rapidly decreasing and a final workingcapacity greater than that of AgX (Ag400B3) was projected. Suchimprovement may be the result of either the base zeolite startingmaterial (13X HP), macropore geometry of the exchanged zeolite and/ordifferences in the final state and/or location of the silver cation inthe zeolite.

Different Ag-exchange levels varying from 10% to 100% were preparedusing the 13XHP base material. CO working capacity was found to beapproximately linearly proportional to Ag-exchange level. Thus, higherAg-exchange level is preferred for the adsorption of CO, with 100%exchange being most preferred.

The commercial AgY and Ag-mordenite adsorbents performed much betterthan conventional zeolites in CO adsorption, but were less effectivethan AgX. The significantly different CO working capacity for theAg-mordenites in Tables II and III are consistent with the higherAg-exchange level of the laboratory prepared Ag-mordenite in Table II,with the lower level of exchanged material (in Table III) having a ΔCOworking capacity of <0.01 mmol/g, and therefore being outside the scopeof the invention.

EXAMPLE 4

In this example, a bed is constructed from a 7.6 cm layer of the AgX ofExample 2 followed by a 15.2 cm layer of palladium supported on porousalumina (Pd/Al₂O₃) oxidation catalyst typical of the prior art. A secondbed was constructed of the same overall length (22.9 cm) using only thePd/Al₂O₃ oxidation catalyst. The oxidation catalyst (E221, 0.5 wt % Pdsupported on the surface of activated alumina beads) was obtained fromDegussa Corporation. Breakthrough tests were performed at the sameconditions as those in Example 3. The results are compared in Table IV.Breakthrough time of H₂ were recorded at 1 ppb, 5 ppb and 20 ppb. TABLEIV tb X_(H2) (H₂) mmol/g hr t_(b) t_(b) @ y_(H2) = y_(H2) = (CO) (CO₂)y_(H2) = Bed y_(H2) = 1 ppb 5 ppb 20 ppb hr hr 20 ppb 7.6 cm AgX + 21.323.2 26.5 >30.7 17.3 0.218 15.2 cm 0.5% PdAl₂O₃ 22.9 cm 3.9 4.913.0 >21.5  3.0 0.072 0.5% PdAl₂O₃

Surprisingly, replacing one third of the catalyst bed with AgX foradsorption of CO resulted in H₂ breakthrough times of two to five timesgreater than for the catalyst used alone. The effective capacity of thecatalyst at 20 ppb H₂ breakthrough was more than tripled when CO wasremoved prior to H₂ oxidation. There was essentially no breakthrough ofCO in either bed over the duration of the test, i.e. 30.7 hours (hr) forthe layered bed and 21.5 hr for the catalyst only bed. Finally, the CO₂breakthrough time (at y_(CO2)=100 ppb) was more than five times greaterwhen CO was adsorbed in AgX. As is evident from the information above,the use of AgX to replace part of the CO/H₂ oxidation catalyst resultsin unexpected advantages in purifying the stream of CO and H₂ whileminimizing CO₂ as an oxidation by-product.

Based upon the examples above, a preferred CO adsorbent for the practiceof the invention has a ΔCO working capacity ≧0.01 mmol/g, preferably≧0.03 mmol/g. In a more preferred embodiment, the CO adsorbent has aΔCO/ΔN₂ separation factor α, as given by Equation 3, equal to or greaterthan 1×10⁻³, preferably greater than 1×10⁻². In the case of small porezeolites (e.g. natural zeolites clinoptilolite, chabazite, etc.), thepore opening or kinetic diameter of the zeolite “window” must be largerthan the kinetic diameter of the CO molecule (0.376 nm). Zeolitesexchanged with Group IB cations (Cu, Ag, Au (gold)) attaining anoxidation state +1 are preferred adsorbents for CO for the applicationof the invention. Ag-exchanged zeolites (>50% exchange) are preferredand highly exchanged AgX (>85% exchange) are most preferred. While notdisclosed above in the examples, gold is believed to be useful in thepractice of the invention given its similarity in chemical structure toCu and Ag. Since the Group IB metals acting as charge balancing cationsin exchanged zeolites provide enhanced CO adsorption capacity, zeolitesrequiring a higher number of charge balancing cations are preferred inthe practice of the invention. Such zeolites are characterized by aSiO₂/Al₂O₃ ratio of <20.

Using the method of this invention, a TSA adsorber and system can bedesigned for the reduction of the concentrations of CO and, optionallyone or more H₂O, CO₂, and H₂ from an incoming feed stream to levels of100 ppb or less, or even 5 ppb or less in the effluent or product gas.Preferably the vessel is used in a TSA prepurifier for an air feedstream. One such vessel design is described below with reference to FIG.3. The arrow (see also in FIGS. 4-6) indicates the direction of gas flowthrough the adsorber bed/vessel. A TSA prepurifier system incorporatingsuch a vessel is disclosed below with reference to FIG. 7.

Returning to FIG. 3, vessel 30 is shown. Vessel 30 contains a firstlayer of H₂O adsorbent (31) such as alumina, silica gel or molecularsieve or mixture of these to remove substantially all of H₂O enteringthe vessel. A second layer (32) of CO₂ adsorbent such as 13X (NaX) or 5Aor mixture of these is used to remove substantially all of CO₂. The CO₂adsorbent layer can also remove any residual water remaining from theH₂O adsorbent layer. A third layer (33) of CO adsorbent is placeddownstream of the CO₂ adsorbent. (By the term “downstream” we meancloser to the effluent or product end of the adsorber vessel.)

A substantially H₂O-free and CO₂-free gas stream enters this COadsorbent layer. The CO adsorbent layer can be designed to remove morethan 50%, preferably more than 95% of the CO in the feed, thus producinga product gas containing <100 ppb CO and most preferably removing morethan 99.8% of the CO in the feed, thus producing a product gas streamcontaining <5 ppb CO. A feed stream, which is substantially free of H₂O,CO₂ and CO, enters the catalyst layer 34. This catalyst removes H₂ andany remaining small amounts of CO using a combination of adsorptionand/or absorption and oxidation. One or more optional adsorbent layers35 may be placed downstream of the CO adsorbent and catalyst layer toremove any CO₂ and H₂O oxidation products formed but not adsorbed in thecatalyst. The optional layer(s) may also be selected to removehydrocarbons, N₂O and/or other trace contaminants.

FIGS. 4-6 show alternative embodiments for integrating a CO adsorbentlayer into an adsorber vessel for use in gas separations such asprepurification. Of course combinations of these embodiments are withinthe scope of the invention.

The required adsorbent and catalyst layer thickness vary according tothe process conditions of experienced by the adsorber during theseparation process. These can vary widely from one system/process toanother. The depth of the layers depicted in FIGS. 3-6 are not intendedto imply or suggest any particular or relative amounts of adsorbentand/or catalyst. The most important process conditions are molar flux ofair, cycle time, feed temperature and pressure and gas composition. Aperson ordinarily skilled in designing a prepurifier should, forexample, be able to design the adsorbent and catalyst layers for eachprepurifier according to the adsorbent/catalyst properties and theprepurifier process conditions. These methods are illustrated in detailin various textbooks such as Ruthven (Principles of Adsorption andAdsorption Processes, 1984).

While FIG. 3 has been described above, FIGS. 4-6 will now be describedbelow.

With reference to FIG. 4, an embodiment is contemplated using vessel 40for the removal of H₂O (in layer 41), CO₂ (in layer 42) and CO (in layer43). In this embodiment a catalyst layer for H₂ removal is not required.

With reference to FIG. 5, a vessel 50 is shown wherein the H₂O and CO₂adsorbents are combined in a single layer 51 either as a singleadsorbent, a mixture of different adsorbents or through the use ofcomposite adsorbents, followed by a CO adsorbent layer 52 and catalystlayer 53 for H₂ removal. A further adsorbent layer (similar in functionto layer 35 illustrated in FIG. 3) for clean-up and/or removal of othercomponents (e.g. hydrocarbons) as with layer 35) may also be used, butis not shown.

With reference to FIG. 6, vessel 60 contains a CO adsorbent layer 62placed prior to the CO₂ adsorbent layer 63, with both being downstreamof H₂O adsorbent layer 61. If removal of H₂ is required an additionalcatalyst layer (not shown) may be used after the CO layer. Additionallayers for clean-up and removal of other components (e.g. hydrocarbons)may be added.

The catalyst used for the removal of H₂ and remaining CO is a supportedmetal catalyst. One or more of the metals Os, Ir, Pt, Ru, Rh, Pd, Fe,Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce may be deposited on a supportchosen from alumina, silica, natural or synthetic zeolites, titaniumdioxide, magnesium oxide or calcium oxide. At least one of these metalsis deposited on the support by techniques well known in the art. Themost preferred catalyst for this invention is Pt and/or Pd supported onalumina.

The optional adsorbent layer described above may be at least one ofalumina, silica gel or zeolite. A zeolite layer is preferred. Mostpreferably, 13X (with 13X APG, available from UOP, Des Plaines, Ill. USAbeing most preferred) is used in this layer when CO₂ and H₂O oxidationproducts are to be removed. This optional H₂O and CO₂ removal layer caneither be placed downstream of the catalyst layer or can be physicallymixed with the catalyst layer. The adsorbent selected for the optionallayer depends upon the contaminants to be removed, e.g. clinoptilolitefor removal of trace N₂O and CO₂ as described in copending commonlyassigned patent application PCT Serial No. O₂/40591 “Method for Removalof N₂O from Gaseous Streams”; and/or NaY, alumina or SELEXORB (acomposite of NaY/alumina available from ALCOA, USA) for the removal ofhydrocarbons.

The invention offers the advantage that existing prepurifiers can beeasily retrofitted by placing the CO adsorbent and optional catalyst atthe downstream end of the prepurifier. The invention may also be appliedto create “high purity” product when ultra high purity product is notrequired, i.e. when the standard purity of existing air separationplants is not sufficient, but electronics grade UHP is not necessary. Insuch applications, any of the embodiments, processes and configurationsof the invention may be used to purify a feed mixture and remove COand/or H₂ contaminants to <100 ppb, individually.

The process is carried out preferably in a cyclic process such aspressure swing adsorption (PSA), temperature swing adsorption (TSA),vacuum swing adsorption (VSA) or a combination of these. The process ofthe invention may be carried out in single or multiple adsorptionvessels operating in a cyclic process that includes at least the stepsof adsorption and regeneration. The adsorption step is carried out atpressure range of 1.0 to 25 bar and preferentially from about 3 to 15bar. The temperature range during the adsorption step is −70° C. to 80°C. When a PSA process is used, the pressure during the regeneration stepis in the range of about 0.20 to 5.0 bar, and preferably 1.0 to 2.0 bar.For a TSA process, regeneration is carried out at a temperature usuallyin the range of about 50° C. to 400° C., preferably between 100° C. to300° C.

One possible process is described herein with reference to FIG. 7. Feedair is compressed in compressor 70 and cooled by chilling means 71 priorto entering one of two adsorbers (76 and 77) where at least thecontaminants H₂O, CO₂ and CO are removed from the air. The adsorbers 76and 77 each have the same adsorbent bed configuration, which may, forexample be one as described with reference to FIGS. 3-6 above. Thepurified air exits the adsorber and then enters the air separation unit(ASU) where it is then cryogenically separated into its major componentsN₂ and O₂. In special designs of the ASU, Ar, Kr and Xe may also beseparated and recovered from the air. While one of the beds is adsorbingthe contaminants from air, the other is being regenerated using purgegas. A dry, contaminant-free purge gas may be supplied from the productor waste stream from the ASU or from an independent source to desorb theadsorbed contaminants and thereby regenerate the adsorber and prepare itfor the next adsorption step in the cycle. The purge gas may be N₂, O₂,a mixture of N₂ and O₂, air or any dry inert gas. In the case of thermalswing adsorption (TSA), the purge gas is first heated in heater 82 priorto being passed through the adsorber in a direction countercurrent tothat of the feed flow in the adsorption step. TSA cycles may alsoinclude a pressure swing. When only pressure swing adsorption (PSA) isutilized, there is no heater.

The operation of a typical TSA cycle is now described in reference toFIG. 7 for one adsorber 76. One skilled in the art will appreciate thatthe other adsorber vessel 77 will operate with the same cycle, only outof phase with the first adsorber in such a manner that purified air iscontinuously available to the ASU. This operation of this out of phasecycle is indicated with reference to the numbers in parentheses.

Feed air is introduced to compressor 70 where it is pressurized. Theheat of compression is removed in chilling means 71, e.g. a mechanicalchiller or a combination of direct contact after-cooler and evaporativecooler. The pressurized, cool and H₂O-saturated feed stream then entersadsorber 76 (77). Valve 72 (73) is open and valves 74 (75), 78 (79) and80 (81) are closed as the adsorber vessel 76 (77) is pressurized. Oncethe adsorption pressure is reached, valve 78 (79) opens and purifiedproduct is directed to an ASU for cryogenic air separation. When theadsorber 76 (77) has completed the adsorption step, valves 78 (79) and72 (73) are closed and valve 74 (75) is opened to blow down the adsorber76 (77) to a lower pressure, typically near ambient pressure. Oncedepressurized, valve 80 (81) is opened and heated purge gas isintroduced into the product end of the adsorber 76 (77). At some timeduring the purge cycle, the heater is turned off so that the purge gascools the adsorber to near the feed temperature.

One of ordinary skill in the art will further appreciate that the abovedescription represents only an example of a typical prepurifier cycle,and there are many variations of such a typical cycle that may be usedwith the present invention. For example, PSA may be used alone whereinboth the heater 82 and the chilling means 71 may be removed.Pressurization may be accomplished with product gas, feed gas or acombination of the two. Bed-to-bed equalization may also be used and ablend step may be incorporated where a freshly regenerated bed isbrought on line in the adsorption step with another adsorber nearingcompletion of its adsorption step. Such a blend step serves to smoothout pressure disturbances due to bed switching and also to minimize anythermal disturbances caused when the regenerated bed is not completelycooled to the feed temperature. Furthermore, the invention may bepracticed with a prepurifier cycle not limited to two adsorber beds. Themethod of the invention can be applied in horizontal, vertical or radialflow vessels.

The method of regeneration depends upon the type of cyclic process. Fora TSA process, regeneration of the adsorbent bed is achieved by passingheated gas countercurrently through the bed. Using the thermal pulsemethod, a cooling purge step follows the hot purge step. The heatedregeneration gas may also be provided at a reduced pressure (relative tothe feed) so that a combined TSA/PSA process is affected. This reducedpressure may be above or below ambient pressure. In cryogenic airseparation processes, the regeneration gas is typically taken from theproduct or waste N₂ or O₂ streams.

In some cases, passing an inert or weakly adsorbed purge gascountercurrently through the bed can further clean the adsorbent bed. Ina PSA process, the purge step usually follows the countercurrentdepressurization step. In case of a single vessel system, the purge gascan be introduced from a storage vessel, while for multiple bed system,purge gas can be obtained from another adsorber that is in theadsorption phase.

The adsorption system can have more steps than the two basic fundamentalsteps of adsorption and desorption. For example, top to top equalizationor bottom to bottom equalization can be used to conserve energy andincrease recovery.

In a preferred embodiment of the invention, as applied to airprepurification, essentially all of water vapor and substantially all ofthe CO₂ are removed from air on at least one layer of activated aluminaor zeolite, or by multiple layers of activated alumina and zeolite priorto passing the air stream through the CO adsorbent layer. Optionally,the CO selective adsorbent layer may be extended and used to remove partor all of the CO₂ from air.

Alternately, in an adsorption vessel, a first layer of adsorbent can beused to remove water vapor and a next layer comprised of a mixture ofthe CO selective adsorbent and 13X (or other zeolite) can be used toremove both CO₂ and CO from the air. Such an adsorbent mixture may becomposed of physically separate adsorbents or of different adsorbentsbound together in the form of a composite.

In applications requiring only a small amount of CO adsorbent (e.g. lowCO concentrations in the feed, etc.), it may be advantageous to mix theCO adsorbent with another adsorbent (such as 13X) in order to affect athicker layer of mixed adsorbent rather than a very thin layer of the COadsorbent alone. The advantage of the thicker mixed layer being ease ofinstallation and less critical tolerance on the overall layer thickness.

The method of the invention can be used to remove one or both of CO andH₂ from air, N₂ and other weakly adsorbing bulk gases. Weaker adsorbinggases to be purified of CO can be identified initially from theelectronic properties of such gases. Using the criteria of CO/N₂selectivity and CO working capacity defined above for selecting anappropriate adsorbent insures that the system will work as well orbetter in removing CO from more weakly adsorbed gases such as N₂, helium(He), neon (Ne), H₂, xenon (Xe), krypton (Kr), argon (Ar), O₂ andmethane (CH₄) and others with similar properties.

Further, while a preferred application of the invention is inprepurification prior to cryogenic air separation, the invention asdescribed herein may also be applied to remove CO and/or hydrogenpresent in nitrogen gas after cryogenic distillation.

The size of H₂O and CO₂ removal layers upstream (by the term “upstream”we mean being located closer to the feed end of the adsorber vessel) ofCO adsorbent would depend upon the H₂O and CO₂ concentration in the gasto be purified. Inert gases such as N₂ may contain ppm levels of O₂, ifthey contain any oxygen at all. Therefore the mechanism for CO and H₂removal in the CO adsorbent and the catalyst would be a combination ofadsorption/chemisorption and absorption. For the removal of CO only byadsorption, an optional clean-up layer may be eliminated as shown inFIG. 4 described above. In the absence of oxygen, the oxidation productsCO₂ and H₂O would not form, therefore an optional clean up layer asshown in FIG. 3 may not be needed. The method of this invention can alsobe used for post purification of N₂ in the cryogenic plants.

As indicated above, the adsorbent beds/vessels used in the method of theinvention can have variety of configurations such as vertical flow beds,horizontal flow beds or radial flow beds and can be operated in apressure swing adsorption mode, temperature swing adsorption mode,vacuum swing adsorption mode or a combination of these.

The adsorbents in this method may be shaped by a series of methods intovarious geometrical forms such as beads, granules and extrudates. Thismight involve addition of a binder to zeolite powder in ways well knownto those skilled in the art. These binders might also be necessary fortailoring the strength of the adsorbents. Binder types and shapingprocedures are well known and the current invention does not put anyconstraints on the type and percentage amount of binders in theadsorbents.

The CO adsorbent could also potentially adsorb some hydrocarbons andnitrogen oxides from air. To ensure complete removal of hydrocarbons,the CO adsorbent can be physically mixed with a hydrocarbon selectiveadsorbent.

The method suggested in this invention can be used for cleanup of anygas containing contaminant levels of CO alone or in combination with H₂,H₂O, CO₂, hydrocarbons and N₂O.

The term “comprising” is used herein as meaning “including but notlimited to”, that is, as specifying the presence of stated features,integers, steps or components as referred to in the claims, but notprecluding the presence or addition of one or more other features,integers, steps, components, or groups thereof.

Specific features of the invention are shown in one or more of thedrawings for convenience only, as each feature may be combined withother features in accordance with the invention. Alternative embodimentswill be recognized by those skilled in the art and are intended to beincluded within the scope of the claims.

1. An adsorption apparatus for the removal of CO from a feed streamcontaining CO in an amount of less than 50 ppm, said apparatuscomprising at least one adsorption vessel containing a CO adsorbentlayer, the CO adsorbent having a ΔCO working capacity greater than orequal to 0.01 mmol/g; and wherein a) when said feed stream furthercontains at least one gas selected from the group consisting ofnitrogen, He, Ne, Ar, Xe, Kr, CH₄ and mixtures thereof, said adsorbentis ion exchanged with a Group IB element; or b) when said feed streamfurther contains at least one gas selected from the group consisting ofoxygen and air and mixtures thereof, said adsorbent is a zeolite havinga SiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with a Ag⁺ or Au⁺. 2.The apparatus of claim 1, wherein said apparatus contains two or more ofsaid adsorption vessels.
 3. The apparatus of claim 1, wherein saidadsorption vessel is selected from the group consisting of vertical flowvessels, horizontal flow vessels, lateral flow vessels or radial flowvessels.
 4. The adsorption apparatus of claim 1, wherein said apparatusfurther contains an adsorbent selective for the adsorption of water, andwherein the water selective adsorbent layer is upstream of said COadsorbent layer.
 5. The adsorption apparatus of claim 1, wherein saidapparatus further contains a catalyst layer for the catalytic oxidationof H₂ to H₂O, and wherein said catalyst layer is downstream of said COadsorbent layer.
 6. The adsorption apparatus of claim 5, wherein saidapparatus further contains an auxiliary adsorbent for the removal ofwater, and wherein said auxiliary adsorbent is downstream of saidcatalyst layer.
 7. The apparatus of claim 1, wherein said ΔCO workingcapacity is greater than or equal to 0.03 mmol/g.
 8. The apparatus ofclaim 4, wherein the water selective adsorbent is one or more of aluminaor NaX zeolite.
 9. The apparatus of claim 5, wherein the H₂ catalyst isa supported metal catalyst.
 10. The apparatus of claim 9, wherein saidsupported metal catalyst is comprises one or more of the metals Os, Ir,Pt, Ru, Rh, Pd, Fe, Co, Ni, Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce and issupported on a substrate selected from the group consisting of alumina,silica, natural or synthetic zeolites, titanium dioxide, magnesium oxideand calcium oxide.
 11. The apparatus of claim 1, wherein said apparatusfurther contains: a) at least one adsorbent layer upstream of said COadsorbent for the adsorption of one or more of H₂O and CO₂, b) acatalyst layer for the catalytic conversion of H₂ to H₂O that isdownstream of said CO adsorbent layer; and c) one or more additionaladsorbents for the adsorption of one or more of H₂O, CO₂, N₂O andhydrocarbons, wherein said additional adsorbents are downstream of saidcatalyst layer.
 12. The apparatus of claim 11, wherein said one or moreadditional adsorbents are selected from the group consisting of alumina,silica gel, clinoptilolite, zeolites, composites thereof and mixturesthereof.
 13. The apparatus of claim 1, wherein the CO adsorbent has aΔCO/AN₂ separation factor of greater than or equal to 1×10⁻³.
 14. Theapparatus of claim 1, wherein when said feed stream further contains atleast one gas selected from the group consisting of nitrogen, He, Ne,Ar, Xe, Kr, H₂, CH₄ and mixtures thereof, said CO adsorbent is selectedfrom the group consisting of AgX zeolite, Ag-Mordenite,Cu-clinoptilolite, AgA zeolite and AgY zeolite.
 15. The apparatus ofclaim 1, wherein when said feed stream further contains at least one gasselected from the group consisting of oxygen and air and mixturesthereof, said CO adsorbent is selected from the group consisting of AgXzeolite, Ag-Mordenite, AgA zeolite and AgY zeolite.
 16. The apparatus ofclaim 1, wherein said CO adsorbent is AgX zeolite.
 17. The apparatus ofclaims 1, wherein when said feed gas further contains air, saidapparatus is an air prepurifier.
 18. The apparatus of claim 1, whereinsaid CO adsorbent is AgX having 100% of its cations associated with Ag.19. A process for the removal for the removal of CO from a feed streamcontaining CO in an amount of less than 50 ppm, said process comprisingcontacting said feed stream with a CO adsorbent having a ΔCO workingcapacity greater than or equal to 0.01 mmol/g to produce a CO depletedgas stream; and wherein a) when said feed stream further contains atleast one gas selected from the group consisting of nitrogen, He, Ne,Ar, Xe, Kr, H₂, CH₄ and mixtures thereof, said adsorbent is a zeoliteexchanged with a Group IB element; or b) when said feed stream furthercontains at least one gas selected from the group consisting of oxygenand air and mixtures thereof, said adsorbent is a zeolite having aSiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with a Ag⁺ or Au⁺.
 20. Theprocess of claim 19, further comprising recovering said CO depleted gasstream, wherein CO is present in said CO depleted gas stream at aconcentration of less than 100 ppb.
 21. The process of claim 19, furthercomprising recovering said CO depleted gas stream, wherein CO is presentin said CO depleted gas stream at a concentration of less than 5 ppb.22. The process of claim 19, wherein said CO concentration in said feedstream is less than 1 ppm CO.
 23. The process of claim 19, wherein saidCO concentration in said feed stream is less than 0.5 ppm CO.
 24. Theprocess of claim 19, wherein said feed gas further comprises water(H₂O), and said process further comprises contacting said feed streamwith a water selective adsorbent that is located upstream of said COadsorbent.
 25. The process of claim 19, wherein said feed gas furthercomprises hydrogen, and said process further comprises contacting saidCO depleted feed stream with a catalyst layer for the catalyticoxidation of H₂ to H₂O to produce a H₂ depleted and H₂O enriched gas,and wherein said catalyst layer is located downstream of said COadsorbent layer.
 26. The process of claim 25, wherein said processfurther comprises the step of contacting said H₂O enriched gas with anadsorbent for the removal of water, and wherein the H₂O adsorbent layeris located downstream of said catalyst layer to produce a gas that isdepleted in CO, H₂ and H₂O.
 27. The process of claim 19, wherein saidΔCO working capacity is greater than or equal to 0.03 mmol/g.
 28. Theprocess of claim 19, wherein the CO adsorbent has aΔCO/AN₂ separationfactor that is greater than or equal to 1×10⁻³.
 29. The process of claim19, wherein the CO adsorbent has a ΔCO/AN₂ separation factor that isgreater than or equal to 1×10⁻².
 30. The process of claim 24, whereinthe water selective adsorbent is one or more of alumina or NaX.
 31. Theprocess of claim 25, wherein the catalyst is a supported metal catalyst.32. The process of claim 31, wherein said supported metal catalyst iscomprises one or more of the metals Os, Ir, Pt, Ru, Rh, Pd, Fe, Co, Ni,Cu, Ag, Au, Zn, Sn, Mn, Cr, Pb, Ce and is supported on a substrateselected from the group consisting of alumina, silica, natural orsynthetic zeolites, titanium dioxide, magnesium oxide and calcium oxide.33. The process of claim 19, wherein said process further comprisespassing said feed gas over: a) at least one adsorbent layer upstream ofsaid CO adsorbent for the adsorption of one or more of H₂O and CO₂, b) acatalyst layer for the catalytic conversion of H₂ to H₂O that isdownstream of said CO adsorbent layer; and c) one or more additionaladsorbents for the adsorption of one or more of H₂O, CO₂, N₂O andhydrocarbons, wherein said additional adsorbents are downstream of saidcatalyst layer.
 34. The process of claim 33, wherein said one or moreadditional adsorbents are selected from the group consisting of alumina,silica gel, clinoptilolite, zeolites, composites thereof and mixturesthereof.
 35. The process of claim 19, wherein said process is selectedfrom the group consisting of pressure swing adsorption, temperatureswing adsorption, or a combination thereof.
 36. The process of claim 19,wherein said process takes place in an adsorber vessel selected from avertical flow vessel, a horizontal flow vessel or a radial flow vessel.37. The process of claim 25, wherein the hydrogen depleted gas containsless than 100 ppb hydrogen.
 38. The process of claim 25, wherein thehydrogen depleted gas contains less than 5 ppb hydrogen.
 39. The processof claim 19, wherein the adsorption step of said process is operated ata temperature of zero to fifty degrees Celsius.
 40. The process of claim19, wherein when said feed gas further contains at least one gasselected from the group consisting of nitrogen, He, Ne, Ar, Xe, Kr, H₂,CH₄ and mixtures thereof said CO adsorbent is selected from the groupconsisting of AgX, Ag-Mor, Cu-clinoptilolite, AgA zeolite and AgYzeolite.
 41. The process of claim 19, wherein when said feed gas furthercontains at least one gas selected from the group consisting of air andoxygen and mixtures thereof said CO adsorbent is selected from the groupconsisting of AgX, Ag-Mor, AgA zeolite and AgY zeolite.
 42. The processof claim 19, wherein said CO adsorbent is AgX having greater than 50% ofits cations associated with Ag.
 43. The process of claim 19, whereinsaid CO adsorbent is AgX having 100% of its cations associated with Ag.44. The process of claim 19, wherein said feed stream contains air, andwherein said CO depleted gas stream is passed to a cryogenicdistillation column.
 45. The process of claim 19, further comprisingrecovering said CO depleted gas stream, wherein CO is present in said COdepleted gas stream at a concentration of less than 1 ppb.
 46. Theprocess of claim 19, wherein the CO partial pressure in said feed streamis less than 0.1 mmHg.
 47. The process of claim 19, wherein the COpartial pressure in said feed stream is less than 0.005 mmHg.
 48. Aprocess for the removal of CO from a feed stream containing CO in anamount of less than 50 ppm and hydrogen said process comprisingcontacting said feed stream with a CO adsorbent having a ΔCO workingcapacity greater than or equal to 0.01 mmol/g to produce a CO depletedgas stream; and wherein said adsorbent is a zeolite exchanged with aGroup IB element.
 49. An adsorption apparatus for the removal of CO froma feed stream containing CO in an amount of less than 50 ppm, andhydrogen, said apparatus comprising at least one adsorption vesselcontaining a CO adsorbent layer, the CO adsorbent having a ΔCO workingcapacity greater than or equal to 0.01 mmol/g; and wherein saidadsorbent is a zeolite exchanged with a Group IB element.
 50. An airprepurification adsorption apparatus for the removal of CO from an airfeed stream containing CO in an amount of less than 50 ppm, saidapparatus comprising at least one adsorption vessel containing a COadsorbent layer, the CO adsorbent having a ΔCO working capacity greaterthan or equal to 0.01 mmol/g; and wherein said adsorbent is a zeolitehaving a SiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with a Ag⁺ orAu⁺.
 51. The apparatus of claim 50, wherein said apparatus furthercomprises at least one adsorbent layer upstream of said CO adsorbent forthe adsorption of one or more of H₂O and CO₂.
 52. The apparatus of claim49 or claim 50, further comprising a catalyst layer for the catalyticconversion of H₂ to H₂O that is downstream of said CO adsorbent layer;and one or more additional adsorbents for the adsorption of one or moreof H₂O, CO₂, N₂O and hydrocarbons, wherein said additional adsorbentsare downstream of said catalyst layer.
 53. The apparatus of claim 49 orclaim 50, wherein said CO adsorbent is AgX.
 54. A process for theremoval for the removal of CO from a feed stream containing CO in anamount of less than 50 ppm and air said process comprising contactingsaid feed stream in an adsorber vessel with a CO adsorbent having a ΔCOworking capacity greater than or equal to 0.01 mmol/g to produce a COdepleted gas stream; and wherein said adsorbent is a zeolite having aSiO₂/Al₂O₃ ratio of <20, and is ion-exchanged with a Ag⁺ or Au⁺.
 55. Theprocess claim 54, wherein said process further comprises passing saidfeed stream over at least one adsorbent layer upstream of said COadsorbent for the adsorption of one or more of H₂O and CO₂ to produce aH₂O and or CO₂ depleted feed stream.
 56. The process of claim 54,wherein said process further comprises, passing said feed stream over acatalyst layer that is downstream of said CO adsorbent layer for thecatalytic conversion of H₂ to H₂O layer and one or more additionaladsorbents for the adsorption of one or more of H₂O, CO₂, N₂O andhydrocarbons, wherein said additional adsorbents are downstream of saidcatalyst layer to produce a feed stream that is depleted in CO, H₂ andone or more of H₂O, CO₂, N₂O and hydrocarbons.
 57. The process of claim54, wherein said CO adsorbent is AgX.
 58. The process of claim 54,wherein the depleted feed stream is passed to a cryogenic distillationcolumn for the separation of air.